Combination process for selected aromatic hydrocarbon production

ABSTRACT

Selected aromatic hydrocarbon concentrates--benzene, mixed xylenes, etc.--are produced by way of a combination process which involves catalytic reforming followed by hot flash separation and dealkylation of the separated hot flash liquid phase. Although the process affords flexibility respecting the precise aromatic concentrate produced, it is particularly directed toward the maximization of benzene.

APPLICABILITY OF INVENTION

As herein described, the present invention relates to a combinationprocess which encompasses catalytic reforming for the conversion ofnaphthenes and paraffins to aromatics, and dealkylation ofalkylaromatics to produce aromatic hydrocarbons having a differentnumber of alkyl substituents. Hydrocarbons classified as aromatics haveenjoyed a continually increasing demand within the market place dueprincipally to their versatility as gasoline blending components and inthe production of a wide spectrum of various petrochemical compounds.Aside from its use as a component of motor fuel, benzene serves as thestarting material for the production of styrene, phenol, syntheticdetergents, DDT and nylon intermediates; principal uses also includefumigants, insecticides and various solvents. The principal use ofethylbenzene, which may be derived by the alkylation of benzene, residesin steam dehydrogenation to produce styrene. Toluene is employed inaviation gasoline and as high-octane blending stocks; as a petrochemicalraw material, it is used in the production of solvents, gums, resins,rubber cement, vinyl organosols and other organic chemicals. Mixedxylenes are primarily used in aviation gasoline and as solvent for alkydresins, lacquers, enamels and rubber cements, etc. Relatively recently,para-xylene has been in great demand for use in the production ofterephthalic acid which is employed in producing synthetic resins andfibers. Cymenes likewise are utilized as solvents and in synthetic resinmanufacture; currently, para-cymene is in great demand for theproduction of para-cresol.

Fresh feed charge stocks for use herein are normally liquid hydrocarbonsboiling within the gasoline, or naphtha boiling range; that is,hydrocarbons which exist in the liquid state at one atmosphere ofpressure and a temperature of about 60° F., and which have normalboiling points up to about 425° F. Thus, it is contemplated thatsuitable charge stocks will include, but not by way of specificlimitation, full boiling range naphthas (about 100° F. to about 400°F.), light naphthas (about 100° F. to about 200° F.) and heavy naphthas(about 20° F. to about 400° F.). As hereinafter indicated in greaterdetail, the charge stock is initially reacted with hydrogen in acatalytic reforming reaction zone, in contact therein with a Group VIIInoble metal catalyst and at reforming conditions which foster theconversion of naphthenes and paraffins to aromatic hydrocarbons.Operating conditions are at a relatively high severity level in order tomaximize the conversion to aromatics. At least a portion, but preferablyall of the resulting catalytically reformed product effluent isseparated in a hot flash zone. The hot flash liquid phase is reacted ina dealkylation reaction zone at conditions which are selected todealkylate alkyl-substituted aromatic hydrocarbons. Product effluentfrom the dealkylation reaction zone is introduced into the hot flashseparation zone, and the intended aromatic concentrate is recovered fromthe resulting hot flash vaporous phase. Alkylaromatics which boil abovethe desired aromatic concentrate are separately recovered and recycledto the dealkylation zone.

In the present specification and the appended claims, the use of theterm "dealkylation" is intended to include both simpledealkylation--e.g. toluene and mixed xylenes are dealkylated tobenzene--and transalkylation where toluene is transalkylated to producea product mixture of benzene and xylenes. As hereinbefore stated, thepresent combination process is primarily intended for the maximumproduction of benzene.

OBJECTS AND EMBODIMENTS

A principal object of my invention is to provide a process for theproduction of a selected aromatic hydrocarbon concentrate. A corollaryobjective resides in a combination process for maximizing the productionof benzene.

A more specific object is directed toward the combination of catalyticreforming and dealkylation, effected in a manner which results in aneconomically enhanced process for aromatic hydrocarbon production andrecovery.

Therefore, one embodiment of the invention herein described is directedtoward a combination process for the production of a selected aromatichydrocarbon concentrate which comprises the sequential steps of: (a)reacting a hydrocarbonaceous charge stock and hydrogen in a catalyticreforming first reaction zone, at reforming conditions selected toconvert paraffins and naphthenes to aromatic hydrocarbons; (b)separating the resulting first reaction zone effluent, in a firstseparation zone, at a temperature not substantially exceeding about 400°F. and a reduced pressure to provide (i) a first vaporous phasecontaining said selected aromatic hydrocarbon concentrate and, (ii) afirst liquid phase containing higher boiling, alkylaromatichydrocarbons; (c) reacting said first liquid phase in a dealkylationsecond reaction zone, at dealkylation conditions selected to dealkylatesaid alkylaromatic hydrocarbons; (d) introducing the resulting secondreaction zone effluent into said first separation zone; (e) separatingsaid first vaporous phase, in a second separation zone, at substantiallythe same pressure and a lower temperature in the range of about 60° F.to about 140° F., to provide (i) a hydrogen-rich second vaporous phaseand, (ii) a second liquid phase containing substantially all of saidaromatic concentrate; (f) recycling at least a portion of saidhydrogen-rich second vaporous phase to said first reaction zone; and,(g) separating said second liquid phase, in a third separation zone, (i)to recover said selected aromatic hydrocarbon concentrate and, (ii) toprovide a concentrated stream of higher boiling aromatic hydrocarbons.

In another embodiment, the concentrated stream of higher boilingaromatic hydrocarbons is recycled for introduction into the dealkylationsecond reaction zone.

These, as well as other objects and embodiments of the presentinvention, will become evident from the following, more detaileddescription thereof. In one such other embodiment, the charge stock is anaphtha fraction having an end boiling point lower than about 400° F.

CITATION OF RELEVANT PRIOR ART

As hereinbefore stated, my inventive concept encompasses the combinationprocess of (1) catalytic reforming for the conversion of naphthenes andparaffins to aromatic hydrocarbons and, (2) dealkylation of theresulting reformed product effluent to produce selected aromatics havinga different number of alkyl substituents. Essentially, the combinationof these two processes encompasses a unique separation technique for therecovery of the desired product. Therefore, no claim is made herein tothe individual processes other than the use thereof in combination withthe product separation technique.

Briefly, the present combination process is effected by initiallyreacting a naphtha boiling range feedstock and hydrogen in a catalyticreforming first reaction zone. Reforming conditions, hereinafterspecifically delineated, are selected to convert the normally liquidhydrocarbons into aromatics--via dehydrogenation anddehydrocyclization--with the concomitant production of hydrogen.Preferably, the entire as-produced reforming zone effluent is introducedinto a hot flash zone, the liquid phase from which serves as the chargeto a dealkylation second reaction zone maintained at dealkylationconditions which convert alkylaromatics to aromatic hydrocarbons havinga lesser number of alkyl substituents. Thus, toluene and mixed xylenesmay be dealkylated for maximum benzene recovery, or toluene may betransalkylated for the purpose of maximizing both benzene and mixedxylenes. In the former situation, a greater quantity of light ends,especially methane, is co-produced, relative to the amount resultingfrom transalkylation.

The dealkylation reaction zone effluent is introduced into the hot flashseparation zone for recovery of additional selected aromatics in the hotflash vapors. Provided, therefore, are (1) a first vaporous phasecontaining the selected aromatic concentrate and some higher boilingaromatic hydrocarbons and, (2) a first liquid phase containing thegreater proportion of higher boiling aromatic hydrocarbons. The latteris recycled to the dealkylation reaction zone while the former isintroduced into a second separation zone at substantially the samepressure and a lower temperature in the range of about 60° F. to about140° F. (commonly referred to as cold, high-pressure separation) toprovide (1) a hydrogen-rich second vaporous phase and, (2) a secondliquid phase containing the selected aromatic concentrate. At least aportion of the hydrogen-rich vaporous phase is recycled to the catalyticreforming reaction zone; a second portion may be recycled to thedealkylation reaction zone. The second liquid phase is separated,preferably by fractional distillation, to remove various light ends andto recover the selected aromatic concentrate, higher boiling aromatichydrocarbons, withdrawn as a bottoms fraction, are preferably recycledto the dealkylation reaction zone.

It is recognized and acknowledged that many illustrations of catalyticreforming, both the traditional fixed-bed system and the more recentstacked system, through which the catalyst particles are downwardlymovable via gravity-flow, are to be found in the prior art. Similarly,the published literature is replete with examples of processes for thedealkylation and/or transalkylation of alkyl-substituted aromatichydrocarbons. Any attempt to delineate exhaustively the appropriateprior art would be an exercise in futility. Therefore, only two examplesof each will be discussed. Since the present technique involves acombination of these two processes, it is believed that the mostrelevant prior art will be directed toward catalytic reforming followedby dealkylation. Copies of the prior art hereinbelow delineatedaccompany this application.

The catalytic reforming section of the present combination process mayutilize a plurality of radialflow, fixed-bed reaction zones, a stackedsystem wherein the catalyst particles are downwardly movable, from onereaction zone to the next succeeding reaction zone, via gravity-flow, ora combination of fixed-bed with the gravity-flowing catalyst system.U.S. Pat. No. 3,706,536 (Cl. 23-288G), issued Dec. 19, 1972, isillustrative of a stacked reactor system in which each reaction zonecontains an annular-form catalyst bed through which the reactant streamflows laterally and radially, while the catalyst particles flowdownwardly from one reaction zone to the next succeeding lower reactionzone. Catalyst particles withdrawn from the last or lowermost zone aretransported to the top of the regeneration tower; regenerated catalystparticles are transported and introduced into the uppermost reactionzone in the stacked system.

U.S. Pat. No. 3,864,240 (Cl. 208-64), issued Feb. 4, 1975, isillustrative of the integration of a reaction system havinggravity-flowing catalyst particles with a fixed-bed system. At leastperiodically, catalyst particles are withdrawn from the former andtransported to the top of a regeneration facility in which they assumethe form of a single descending column. Regenerated catalyst particlesare subsequently transported to the top of the gravity-flowing reactionzone.

As those possessing the requisite skill in the appropriate art areaware, the described continuous catalyst regeneration reforming systemoffers numerous advantages when compared to the conventional fixed-bedprior art systems. Among these is the capability of efficient operationat comparatively lower pressures and higher liquid hourly spacevelocities. With continuous catalyst regeneration, higher consistentinlet catalyst bed temperatures can be maintained, and there is acorresponding increase in both hydrogen production and the puritythereof within the recycled vaporous phase.

Dealkylation of alkylaromatic hydrocarbons, in a catalytic system, isdisclosed in U.S. Pat. No. 3,197,523 (Cl. 260-672), issued July 27,1965. Suitable alkylaromatic feedstocks are those comprising toluene,mixed xylenes, ethylbenzene, mixed diethylbenzenes and variousalkyl-substituted naphthalenes. Catalysts utilized in the processcontain at least one oxide of tin, titanium and zirconium combined withat least one oxide of chromium, molybdenum and tungsten. Disclosedoperating conditions include temperatures in the range of about 1000° F.to about 1500° F. and pressures from about 300 psig. to about 1000 psig.

Another hydrodealkylation technique is described in U.S. Pat. No.3,204,007 (Cl. 260-672), issued Aug. 31, 1965; the process isparticularly directed toward the production and recovery ofnon-substituted aromatic hydrocarbons, principally benzene and/ornaphthalene. Hydrodealkylated product effluent is cooled by beinginitially utilized as a heat-exchange medium (to preheat fresh feed tothe direct-fired heater) and then passed into a cooler/condenser (9).Cooled effluent traverses three flash drums, the liquid phase passingtherethrough in series. The three flash drums function at succeedinglylower pressures of 500 psig. to 600 psig. (12), 50 psig. to 150 psig.(22) and atmospheric pressures (31). In view of the intentional coolingand condensing of the dealkylated product effluent, these three zonesare considered "cold flash separators" as distinguished from a "hotflash zone".

Operating conditions include temperatures from about 1000° F. to 1500°F., pressures in the range of about 300 psig. to about 600 psig., aliquid hourly space velocity of 0.1 to 20.0 (preferably 0.5 to 5.0) anda hydrogen to hydrocarbon fresh feed mole ratio of about 8.7:1.0 (referto Example 1). Catalysts comprise metals from the platinum group,cesium, tungsten, silver, rhenium and chromium combined with a highsurface area carrier material.

Transalkylation of alkylaromatic hydrocarbons constitutes the subject ofU.S. Pat. No. 3,763,260 (Cl. 260-672T), issued Oct. 2, 1973. Here thecatalytic composite constitutes a metal component selected from thegroup of copper, silver and zirconium combined with zeolitic mordenitehaving a silica/alumina mole ratio of at least 40.0:1.0. Such catalystsare utilized for hydrocarbon transalkylation reactions at temperaturesfrom 0° C. to about 500° C. (32° F. to about 932° F.) and pressures fromabout atmospheric to about 1500 psig. Applicable reactions includetransalkylation of toluene to produce benzene and mixed xylenes,transalkylation of toluene with C₉ -methyl aromatics to produce xylenesand transalkylation of benzene with polyethylbenzene to produceethylbenzene. U.S. Pat. No. 3,780,122 (Cl. 260-672T), issued Dec. 18,1973 is believed to be cumulative to the foregoing, the exception beinga lack of disclosure of metallic components being combined with thezeolitic mordenite.

A combination process of catalytic reforming and dealkylation ispresented in U.S. Pat. No. 3,371,126 (Cl. 260-672), issued Feb. 27,1968. This is additionally combined with both a hydrogen-producing plant(via steam reforming of naphtha) and a hydrogen purification system (viacryogenic techniques). Principally, the objective of this processresides in the simultaneous production of benzene-type hydrocarbons andtown gas of a predetermined calorific value. The catalytic reforming anddealkylation reaction zones function in series with the total effluentfrom the former being directly introduced into the latter. Dealkylationreaction product effluent is withdrawn and condensed (column 6, lines70-71), and introduced into a cold separator (24) for separation intonormally vaporous components and normally liquid components from whichbenzene is recovered via fractionation. Clearly, there is a lack ofrecognition respecting the use of a hot flash zone to separate thereforming zone product effluent. Further, there exists no teaching ofintroducing the hot flash vaporous phase into a cold separator.

It is believed that the foregoing, taken either singularly or incombination, neither anticipates, nor renders the present invention, asdescribed and claimed herein, obvious.

SUMMARY OF INVENTION

Essentially, the combination process encompassed by my inventive conceptinvolves two reaction systems, catalytic reforming and dealkylation (ortransalkylation) and a three-step separation facility for the ultimatelydesired product recovery. The separation facility consists of a hotflash zone, a cold separator into which condensed hot flash vapors areintroduced and a fractionation facility which recovers the desiredproduct from the cold separator liquid phase.

As hereinbefore stated, the catalytic reforming, system may functionwith a plurality of fixed-bed zones, with a plurality of stacked zonesthrough which catalyst particles flow via gravity, or a combinationthereof as described in the previously discussed U.S. Pat. No.3,864,240. Since aromatic hydrocarbon production, via dehydrogenation ofnaphthenes and especially dehydrocyclization of paraffins, is favored bya relatively high-severity operation--e.g. higher temperatures and lowerpressures--the continuous catalyst regeneration reforming system isparticularly preferred. This preference stems from the fact that ahigher carbon (often referred to as "coke") level can be tolerated onthe catalytic composites. Similarly, a mixed system wherein the naphthacharge stock is first serially reformed in two or more fixed-bed zones,followed by one or more zones in the stacked, gravity-flowingconfiguration, may be utilized. The latter has the further advantage inaffording the utilization of two different catalysts which permits theoverall reforming process to be "tailored" to achieve the desiredreformed product.

Catalytic reforming of naphtha boiling range hydrocarbons is avapor-phase operation, and is effected at conversion conditionsincluding catalyst bed temperatures in the range of about 750° F. toabout 1020° F.; judicious and cautious techniques generally dictate thatcatalyst temperatures not substantially exceed a level of about 1020° F.Other conditions include pressures in the range of about 50 psig. toabout 1000° psig., a liquid hourly space velocity (defined as volumes offresh charge stock per hour, per volume of total catalyst particles) inthe range of about 0.5 to about 10.0 and a hydrogen to hydrocarbon moleratio in the range of about 1.0:1.0 to about 15.0:1.0. As a practicalmatter, fixed-bed reforming systems necessitate lower catalyst bedtemperatures from 750 ° F. to about 910° F., higher pressures from about500 psig. to 1000 psig., lower space velocities of 0.5 to about 2.5 andhigher hydrogen/hydrocarbon mole ratios of 4.5:1.0 to about 8.0:1.0. Onthe other hand, benefits accrue through continuous catalyst regenerationreforming in that the operating conditions involve higher catalyst bedtemperatures from 950° F. to 1010° F., lower pressures of about 50 psig.to about 450 psig., higher space velocities of 3.0 to about 8.0 andlower hydrogen/hydrocarbon mole ratios of 0.5:1.0 to about 5.5:1.0.

Catalytic reforming reactions are varied, and include dehydrogenation ofnaphthenes to aromatics, dehydrocyclization of paraffins to aromatics,hydrocracking of long-chain paraffins into lower-boiling, normallyliquid material and, to a certain extent, the isomerization ofparaffins. These reactions, the net result of which is endothermicitywith respect to the overall reaction system, are effected through theutilization of one or more Group VIII noble metals--e.g. platinum,palladium, rhodium, ruthenium, osmium and iridium--combined with ahalogen, generally chlorine and/or fluorine, and a porous carriermaterial such as gamma alumina. Relatively recent investigativedevelopments have indicated that unexpected advantageous results areattainable and enjoyed through the cojoint utilization of a catalyticmodifier; these have been selected from the group of cobalt, nickel,gallium, germanium, tin, rhenium, vanadium, tungsten, zinc and mixturesthereof.

The precise operating conditions and catalytic composite utilized in thecatalytic reforming section will be dependent upon both the physical andchemical characteristics of the naphtha boiling range charge stock, aswell as upon the selected aromatic concentrate to be recovered.Therefore, it is understood that the viability of the present inventiondoes not rely upon either catalyst composition, or operating conditionsin the reforming section. The naphtha feedstock may be, and mostgenerally is pre-heated via indirect heat-exchange with one or morehigh-temperature process streams. For the most part, such streams willinclude the effluent withdrawn from both the reforming and dealkylationsections; the latter will normally have the higher temperature. In anyevent, the charge stock is subsequently introduced into a direct-firedheater wherein its temperature is further increased to at least thelevel needed to provide the designed temperature at the inlet to thecatalyst bed and the initial reaction zone. Since reforming reactionsare generally endothermic, and reforming is effected in a plurality ofindividual zones, the temperature of the effluent from one zone will beincreased in an interstage heater prior to passing into the nextsucceeding reaction zone.

In accordance with the present combination process, the catalyticallyreformed product effluent, preferably in total, is introduced into thehot flash separation zone at a temperature not substantially in excessof about 400° F. and at a reduced pressure. As above stated, the producteffluent from the catalytic reforming section may be first employed as apre-heat medium. Another such use would be to supply at least a portionof the heat necessary to effect the final fractional distillation forproduct recovery. Likewise, the dealkylation reaction zone effluent mayserve to supply heat to the fresh feed charge stock, or to the feed tothe fractionation facilities, either in and of itself, or in some "hotoil belt" combination with the reformed product effluent. The purpose ofthe hot flash zone is to provide (i) a vaporous phase containing all theselected aromatic concentrate first produced in the catalytic reformingsystem and, (ii) a normally liquid phase containing higher boilingaromatic hydrocarbons. Normally vaporous hydrocarbons, hydrogen andnormally liquid material boiling below the selected aromatichydrocarbons are also removed in the vaporous phase. The reducedpressure under which the hot flash zone functions is dependent upon thetemperature at which the feed thereto is introduced. Considered also isthe character of the aromatic hydrocarbons which are intended to beflashed and recovered in the vaporous phase.

The liquid phase from the hot flash separation zone is introduced intothe dealkylation reaction zone in order to effect further conversion ofalkylaromatics to the selected aromatic concentrate. Dealkylationoperating conditions will generally be within the ranges suggested bythe prior art previously delineated. As stated, these includetemperatures within the range of about 1000 ° F. to about 1500° F.(believed to refer to the temperature at the inlet to the catalyst bed),pressures from about 300 psig. to about 1000 psig. and liquid hourlyspace velocities preferably in the range of about 0.5 to about 5.0. Whenutilizing zeolitic-based catalytic material, lower temperatures fortransalkylation are afforded; in U.S. Pat. Nos. 3,763,260 and 3,780,122(previously discussed), temperatures of 500° F. to 800° F. and 392° F.to 896° F. are respectively disclosed.

In contrast to catalytic reforming, which is an endothermic,hydrogen-producing reaction, dealkylation is an exothermic,hydrogen-consuming reaction. Consequently, the temperature of thereactant stream effluent exiting the dealkylation reaction zone will beconsiderably higher than the temperature at which the catalyticallyreformed product effluent is withdrawn from the reforming section. Inthose cases where the selected aromatic concentrate requirestransalkylation to produce benzene and mixed xylenes from toluene, theexit temperature will be lower than that experienced when dealkylationis being effected to maximize benzene production. Regardless, theexothermic dealkylation, or transalkylation reactions may be effected ina plurality of reaction zones having intermediate cooling facilities inorder to decrease the overall temperature differential. Such a techniqueis well known in the art of conducting exothermic reactions, and formsno essential feature of the present invention.

Suitable dealkylation and transalkylation catalysts, for utilization inthe present combination process, include those alluded to in theprevious discussion of the prior art. Such catalysts generally consistof one or more catalytically active metallic components combined with asuitable inorganic oxide carrier material. In many instances, the acidicfunction of the catalyst will be enhanced through the addition theretoof a halogen component, particularly a chlorine and/or fluorinecomponent. Inorganic oxide carrier materials include both amorphous andzeolitic material, as well as mixtures thereof. When amorphous, thecarrier is generally selected from the group of alumina, silica,zirconia, titania, hafnia, boria and various mixtures thereof. Zeoliticcarriers are of the character of crystalline aluminosilicates, andinclude mordenite, Type X and Type Y molecular sieves; such zeoliticmaterial may be incorporated within an amorphous matrix.

With respect to the metallic components, the selections are availablefrom a large number. These too may be those of the prior art includingtin, titanium, zirconium, chromium, molybdenum, tungsten, silver,copper, rhenium, and the noble metals of Group VIII, such as platinum,osmium, rhodium, ruthenium, iridium and palladium. From the practicalviewpoint, the selection of the particular catalytic composite will bedependent upon the component analysis of the aromatic hydrocarbonportion of the reformed product effluent which is charged to thedealkylation reaction zone. Preferred for use herein, when the principalreactions involve dealkylation for maximum benzene, is gamma, or etaalumina containing a chromium component; for the transalkylation oftoluene, to produce benzene and mixed xylenes, the use of zeoliticmordenite, which may be admixed within an amorphous matrix is preferred.

Vaporous material from the hot flash zone is cooled and condensed to atemperature generally in the range of about 60° F. to about 140° F., andintroduced at substantially the same pressure into a cold separator.Normally gaseous hydrocarbons and hydrogen are removed as a vaporousphase; in some instances, this stream will contain a minor quantity ofpentanes and hexanes. A portion of the gaseous material is vented fromthe system under pressure control, and at least another portion isrecycled to the catalytic reforming system. Flexibility is afforded inthat another portion can be diverted to the dealkylation reaction zone.The normally liquid phase from the cold separator will contain pentanesand heavier hydrocarbons, including substantially all of the desiredaromatics, and some absorbed gaseous paraffins and hydrogen. The latterserves as a feed to a fractionation facility, from which the selectedaromatic hydrocarbon fraction is removed as a heart-cut, lower boilingcomponents are withdrawn as an overhead stream and the higher boilingaromatics as a bottoms fraction. In a preferred technique, at least aportion of this bottoms fraction is recycled to the dealkylationreaction zone. Alternatively, all of the bottoms fraction may be sorecycled, or all may be withdrawn as a separate product of the process.

BRIEF DESCRIPTION OF DRAWING

Further description of the process encompassed by my inventive conceptwill be made in conjunction with the accompanying drawing, which ispresented for the sole purpose of illustration, and not with the intentof limiting the invention beyond the scope and spirit of the appendedclaims. The drawing is shown as a simplified schematic flow diagram inwhich details such as pumps, instrumentation and other controls,coolers, condensers, compressors, heat-exchange and heat-recoverycircuits, valving start-up lines and similar hardware have beeneliminated or reduced in number as being non-essential to theunderstanding of the techniques involved. Utilization of thesemiscellaneous appurtenances, to modify the process as illustrated, willbe evident to those possessing the requisite skill in the art ofpetroleum refining technology.

In the drawing, reforming reactor 5 will be described as athree-reaction zone, stacked system through which catalyst particles aremovable via gravity-flow. The initial direct-fired charge heater and twocatalyst bed inter-heaters are not illustrated. Similarly, reactor 10constitutes the dealkylation reaction zone, and will be described as twoseries-flow chambers having cooling facilities therebetween.

DETAILED DESCRIPTION OF DRAWING

With specific reference now to the drawing, fresh feed charge stock,being a normally liquid naphtha fraction having an initial boiling pointof about 180° F. and an end boiling point of about 390° F., isintroduced into the process by way of conduit 1. This heavy naphthafraction has been previously subjected to hydrotreating for olefinichydrocarbon saturation, and for the removal of sulfurous and nitrogenouscomponents. The charge stock is initially introduced into heat-exchanger2 wherein it is preheated via indirect contact with reformed producteffluent from line 3. Continuing through line 1, the fresh feed isadmixed with a hydrogen-rich gaseous phase in line 4, in the amount suchthat the hydrogen/fresh feed mole ratio is about 2.5:1.0. The mixturecontinues through line 4, and is introduced thereby into reformingreactor system 5, after being increased in temperature to about 1000° F.As above set forth, the catalytic reforming is preferably effected in astacked system, in which catalyst particles are downwardly-movable byway of gravity-flow, integrated with a continuous catalyst regenerationtower.

In this illustration, the stacked system consists of three individualreaction zones containing a catalytic composite of about 0.6% platinum,0.5% tin and about 1.0% chlorine, by weight and calculated as theelements thereof; these catalytically active ingredients are combinedwith gamma alumina. The pressure at the inlet to the stacked system isabout 420 psig., while at the outlet, the pressure is about 410 psig.The inlet temperature to each of the three beds of catalyst particles ismaintained at about 990° F., and the overall liquid hourly spacevelocity approximates 1.6. Catalytically reformed effluent is withdrawnby way of conduit 3 and is employed as the heat-exchange medium inexchanger 2. The cooled effluent is withdrawn through line 6 and admixedwith dealkylation product effluent from line 11. The mixture continuesthrough line 6, and is introduced thereby into hot flash separation zone7 at a temperature in the range of about 250° F. to about 350° F. and apressure of about 200 psig.

In this particular situation, the intent is to maximize the productionof benzene from the heavy naphtha charge stock. Therefore, hot flashseparation zone 7 will function to produce (i) a vaporous phase in line12 which contains substantially all the benzene and lower-boilinghydrocarbons and, (ii) a liquid phase in line 8 which is combined with aportion of a hydrogen-rich gaseous phase in line 9, the mixturecontinuing through line 9 as the feed to dealkylation reaction zone 10.The catalytically reformed product effluent portion of the total feed tohot flash zone 7, excluding pentanes, normally vaporous hydrocarbons andhydrogen, has the approximate component stream analysis shown in thefollowing Table I:

                  TABLE I                                                         ______________________________________                                        Normally Liquid Reformed Effluent                                             Component        Bbl./day    Vol. %                                           ______________________________________                                        Hexane           264         2.33                                             C.sub.7 --C.sub.9 Paraffins                                                                    1047        9.20                                             Benzene          466         4.10                                             Toluene          1309        11.51                                            Mixed Xylenes    2663        23.41                                            Ethylbenzene     597         5.25                                             C.sub.9 --plus Aromatics                                                                       5028        44.20                                            ______________________________________                                    

Dealkylation reaction system 7 contains a catalytic composite of 1.9% byweight of titanium oxide, 4.7% of tin oxide and about 15.0% of chromiumoxide, combined with gamma alumina. In this illustration, thedealkylation reactions are effected at a catalyst bed inlet temperatureof 1,200° F. and a pressure of about 400 psig. Exothermicity results ina catalyst bed outlet temperature of about 1,350° F.; therefore in thetwo-zone series system, the first zone effluent will be cooled to about1,200 ° F. prior to being introduced into the second reaction zone. Aspreviously set forth, hot flash separation zone serves to removesubstantially all the benzene and lighter hydrocarbons as a vaporousphase in line 12. This phase will also contain an amount of the heavieraromatic hydrocarbons, much the same as the liquid phase in line 8 willcontain some of the benzene and hexane. For the purposes of thisillustration, it will be presumed that the hot flash zone liquid phaseis substantially completely free from material boiling below toluene.

The hot flash liquid phase in line 8 is admixed with a hydrogen-richgaseous phase in line 9, the mixture continuing therethrough intodealkylation reaction system 10. Dealkylation product effluent, at atemperature of about 1,350° F., is withdrawn via line 11; after its useas a heat-exchange medium and further cooling to a temperature of about250° F. to about 350° F., the effluent is admixed with the reformedproduct effluent in line 6, and introduced therewith into hot flash zone7. The vaporous phase in line 12, which now is inclusive of the benzeneand lower-boiling material produced via dealkylation, is introduced intocooler/condenser 13, wherein the temperature is lowered to a level inthe range of 60° F. to about 140° F.--e.g. 90° F. The thus-condensedvaporous phase passes through conduit 14 into cold separator 15 at apressure of about 190 psig.

A hydrogen-rich vaporous phase is recovered via line 4, and a portionthereof is vented from the process by way of conduit 16. The remaindermay be subjected to cryogenic separation in order to increase thehydrogen purity prior to being recycled to the catalytic reformingsystem. Preferably, a second portion of the recycled hydrogen-rich phaseis diverted via line 9 for introduction into dealkylation system 10. Theamount so diverted should be such that hydrogen is present in excess ofthat which will be consumed during the dealkylation reactions. Condensedliquid material is introduced through conduit 17 into a fractionationfacility 18. Fractional distillation conditions of temperature andpressure are such that hexanes and lower-boiling hydrocarbons arewithdrawn as an overhead stream through conduit 19; benzene concentrateis withdrawn as a heart-cut in line 20. Heavier aromatic hydrocarbonsare recovered as a bottoms stream through line 21, and recycled therebyinto dealkylation reactor 10. Where either desirable, or necessary toprevent the buildup of a refractory component in the recycled material,a drag stream may be withdrawn from the process through line 22.

In the following Table II, the material balance respecting only thedealkylation system is presented; the numerical values representthousands of pounds per day. That is, the indicated yields do notinclude the 466 Bbl./day of benzene originally part of the reformedproduct effluent recovered in the vaporous phase from hot flash zone 10.Likewise, the tabulation does not include any material withdrawn orvented from the process through conduits 16, 19 and which was originallyin the reformed effluent in line 3, and/or conduit 22.

                  TABLE II                                                        ______________________________________                                        Dealkylation Material Balance                                                                     Reactor                                                   Component  Feed     Effluent  Recycle                                                                              Net                                      ______________________________________                                        Hydrogen   153.3*   --        --     --                                       Methane    --       1238.2    --     1238.2                                   Ethane     --       12.8      --     12.8                                     Benzene    --       2001.8    --     2001.8                                   Toluene    399.5    326.2     326.2  --                                       C.sub.8 --Aromatics                                                                      996.2    70.2      70.2   --                                       C.sub.9 --Aromatics                                                                      1443.6   6.0       6.0    --                                       C.sub.7 --C.sub.9 Paraffins                                                              260.2    --        --     --                                       ______________________________________                                         *Hydrogen consumed in dealkylation reactions.                            

The foregoing specification, particularly when read in conjunction withthe accompanying drawing, is believed to present a clear understandingof the present invention, the scope of which is defined by the appendedclaims.

I claim as my invention:
 1. A combination process for the production ofa selected aromatic hydrocarbon concentrate which comprises thesequential steps of:(a) reacting a hydrocarbonaceous charge stock andhydrogen in a catalytic reforming first reaction zone, at reformingconditions selected to convert paraffins and naphthenes to aromatichydrocarbons; (b) separating the resulting first reaction zone effluent,in a first separation zone, at a temperature of at least 250° F. but notsubstantially exceeding about 400° F. and a reduced pressure to provide(i) a first vaporous phase containing said selected aromatic hydrocarbonconcentrate and, (ii) a first liquid phase containing higher boiling,alkylaromatic hydrocarbons; (c) reacting said first liquid phase in adealkylation second reaction zone, at dealkylation conditions selectedto dealkylate said alkylaromatic hydrocarbons; (d) introducing theresulting second reaction zone effluent into said first separation zone;(e) separating said first vaporous phase, in a second separation zone,at substantially the same pressure and a lower temperature in the rangeof about 60° F. to about 140° F., to provide (i) a hydrogen-rich secondvaporous phase and, (ii) a second liquid phase containing substantiallyall of said aromatic concentrate; (f) recycling at least a portion ofsaid hydrogen-rich second vaporous phase to said first reaction zone;and, (g) separating said second liquid phase, in a third separationzone, (i) to recover said selected aromatic concentrate and, (ii) toprovide a concentrated stream of higher boiling aromatic hydrocarbons.2. The process of claim 1 further characterized in that said higherboiling aromatic hydrocarbons are introduced into said second reactionzone.
 3. The process of claim 2 further characterized in that at least aportion of said hydrogen-rich second vaporous phase is introduced intosaid second reaction zone.
 4. The process of claim 1 furthercharacterized in that said hydrocarbonaceous charge stock consists ofnormally liquid hydrocarbons boiling up to about 425° F.
 5. The processof claim 1 further characterized in that said charge stock is a naphthahaving an end boiling point lower than about 400° F.
 6. The process ofclaim 1 further characterized in that said selected aromatic hydrocarbonconcentrate is benzene.
 7. The process of claim 6 further characterizedin that said concentrated stream of higher boiling aromatic hydrocarbonsis introduced into said second reaction zone.
 8. The process of claim 1further characterized in that said selected aromatic concentratecomprises benzene, toluene and xylene.
 9. The process of claim 8 furthercharacterized in that said concentrated stream of higher boilingaromatic hydrocarbons is introduced into said second reaction zone. 10.The process of claim 1 further characterized in that said selectedaromatic concentrate is benzene, a first portion of said higher boilingaromatic hydrocarbons is separately recovered and a second portion isintroduced into said second reaction zone.